Integrated production of hydrogen, electricity, and heat

ABSTRACT

A method and a system for the coproduction of hydrogen, electrical power, and heat energy. An exemplary method includes desulfurizing a feed stream to form a desulfurized feed stream, reforming the desulfurized feed stream to form a methane rich gas, and providing the methane rich gas to a membrane separator. A hydrogen stream is produced in a permeate from the membrane separator. A retentate stream from the membrane separator is provided to a solid oxide fuel cell (SOFC). Electrical power is produced in the SOFC from the retentate stream.

BACKGROUND

A steadily increasing amount of carbon dioxide (CO₂) in the atmospherehas intensified environmental challenges such as climate change andocean acidification around the world. CO2 alone accounts for over 75% ofall greenhouse gas emissions. In order to slow down the potentialnegative impact of CO2 on the environment scientific studies arefocusing on capturing, converting and sequestering CO2 from sectors thataccount for most of the emissions such as transport and powergeneration.

Electrification of transport sector is happening at rapid pace withemergence of hydrogen fuel cell electric vehicles and battery electricvehicles, which allows elimination of point sources of emissions such aspassenger and heavy-duty vehicles. Challenges remain specifically inmeeting infrastructure requirements for supplying hydrogen andelectricity for these new applications. An efficient global hydrocarbondistribution infrastructure already exists in many parts of the worldand this invention allows the economic utilization of thisinfrastructure in addressing new market demands for energy while thehydrogen and electricity infrastructure is established.

SUMMARY

An embodiment described herein provides a method for coproduction ofhydrogen, electrical power, and heat energy. The method includesdesulfurizing a feed stream to form a desulfurized feed stream,reforming the desulfurized feed stream to form a methane rich gas, andproviding the methane rich gas to a membrane separator. A hydrogenstream is produced in a permeate from the membrane separator. Aretentate stream from the membrane separator is provided to a solidoxide fuel cell (SOFC). Electrical power is produced in the SOFC fromthe retentate stream.

Another embodiment described herein provides a trigeneration facility.The trigeneration facility includes a desulfurization unit to removesulfur from a hydrocarbon feed stream forming a desulfurized feedstream. A pre-reformer is included to convert the desulfurized feedstream to a methane rich gas. A membrane separator is included to removeat least a portion of hydrogen from the methane rich gas in a permeate.A solid oxide fuel cell (SOFC) is included to generate electrical powerfrom a retentate from the membrane separator.

BRIEF DESCRIPTION OF DRAWINGS

FIG. 1 is a schematic drawing of a trigeneration facility thatincorporates a hydrogen membrane separator.

FIG. 2 is a schematic drawing of a trigeneration facility thatincorporates a hydrogen membrane separator integrated with a water-gasshift reactor.

FIG. 3 is a flowchart of a process for using a hydrocarbon feed streamto produce hydrogen, electricity, and heat in a trigeneration facilitythat incorporates a hydrogen membrane separator.

FIG. 4 is a flowchart of a process for using a hydrocarbon feed streamto produce hydrogen, electricity, and heat in a trigeneration facilitythat incorporates a hydrogen membrane separator integrated with awater-gas shift reactor.

FIG. 5 is a simplified process model of a trigeneration facility thatincorporates a hydrogen membrane separator.

FIG. 6 is a simplified process model of a trigeneration facility thatincorporates a hydrogen membrane separator integrated with a water-gasshift reactor.

DETAILED DESCRIPTION

The systems described herein address technical, system andinfrastructure challenges faced by transport industry that aspires toachieve enhanced efficiency and sustainability. Examples describedherein provide a trigeneration facility that includes a desulfurizationunit, a pre-reformer, a membrane separator, and a solid oxide fuel cell(SOFC). In the desulfurization unit, sulfur and other impurities areremoved from a hydrocarbon feed stream generally having 12 carbon atomsor fewer, for example, with a boiling point of less than about 200° C.The resulting desulfurized stream is treated in a pre-reformer togenerate a methane rich gas stream that contains hydrogen.

The reformate from the pre-reformer flows through a membrane separatorwhich removes hydrogen from the reformate. The hydrogen may be used as azero carbon product for transportation, power generation, and otheruses, such as in a chemical plant. A portion of the hydrogen may be usedin the desulfurization unit, for example, for hydrodesulfurization.After the removal of at least a portion of the hydrogen the remainingreformate is fed to the SOFC for the generation of electrical power.Heat energy generated in the SOFC may be used in other processes, suchas steam generation and the like. At least a portion of the heat energymay be used to heat the remaining reformate stream to the operatingtemperature of the SOFC prior to introducing it to the SOFC. A wastestream from the SOFC, including CO2, can be purified for use in CO2conversion processes or sequestration.

In the facility, the hydrocarbon feed stream is directly processed togenerate electrical power, hydrogen, and heat. In addition to power,these technologies can be used to produce synthetic gases andpotentially hydrogen. When targeting low carbon footprint operations,the CO₂ in the waste stream is captured, used, or stored.

FIG. 1 is a schematic drawing of a trigeneration facility 100 thatincorporates a hydrogen membrane separator 102. The trigenerationfacility 100 has a hydrocarbon feed 104 that can include any number oflight hydrocarbons. In various embodiments, the hydrocarbon feed 104incudes natural gas, liquefied petroleum gas (city gas), or naphtha witha boiling point range up to 200 degree Celsius.

In the embodiments of FIG. 1, a desulfurization unit 106 removes sulfurand other impurities, for example, in a hydrodesulfurization orhydrodemetallization process. In some embodiments, the desulfurizationis performed by adsorbents at ambient temperature. These embodiments mayprovide simpler operation and maintenance when the amount of sulfur inthe gas is relatively low, e.g., less than about 1%, less than about0.5%, less than about 1500 ppm, less than about 500 ppm, or lower.

In embodiments using a desulfurization catalyst, a hydrogen stream maybe provided to the desulfurization unit 106. In the desulfurization unit106, contaminants such as metals, sulfur, and nitrogen can be removed bypassing the hydrocarbon feed 104 through a series of layered catalyststhat perform the catalytic functions of one or more of demetallization,desulfurization, and denitrogenation. In some embodiments, the sequenceof catalysts to perform hydrodemetallization (HDM) andhydrodesulfurization (HDS) can include a hydrodemetallization catalyst,an intermediate catalyst, a hydrodesulfurization catalyst, and a finalcatalyst.

An intermediate catalyst can be used to perform a transition between thehydrodemetallization and hydrodesulfurization functions. Theintermediate catalyst can have intermediate metal loadings and pore sizedistribution. The catalyst in the desulfurization unit 106 can be analumina based support in the form of extrudates, at least one catalyticmetal from group VI (for instance, molybdenum, tungsten, or both), or atleast one catalytic metals from group VIII (for instance, nickel,cobalt, or both), or a combination of any two or more of them. Thecatalyst can contain at least one dopant, such as one or more of boron,phosphorous, halogens, and silicon. The intermediate catalyst can have asurface area of between about 140 m²/g and about 200 m²/g, a pore volumeof at least about 0.6 cm³/g, and mesoporous pores sized between about 12nm and about 50 nm.

The catalyst to perform the desulfurization can include gamma aluminabased support materials with a surface area towards the higher end ofthe HDM range, such as between about 180 m²/g and about 240 m²/g. Thehigher surface for the HDS catalyst results in relatively smaller porevolume, such as a pore volume of less than about 1 cm³/g. The catalystcontains at least one element from group VI, such as molybdenum, and atleast one element from group VIII, such as nickel. The catalyst alsocontains at least one dopant, such as one or more of boron, phosphorous,silicon, and halogens. In some examples, cobalt (Co) can be used toprovide relatively higher levels of desulfurization. The metals loadingfor the active phase is higher as the desired activity is higher, suchthat the molar ratio of Ni:(Ni+Mo) is between about 0.1 and about 0.3and the molar ratio of (Co+Ni):Mo is between about 0.25 and about 0.85.

As described herein, the desulfurization unit 106 processes thehydrocarbon feed 104 with hydrogen, for example, from the hydrogenmembrane separator 102. The hydrogen may be added at 0.1 mol. %, 0.5mol. %, 1 mol. %, 5 mol. %, or higher, as a proportion of thehydrocarbon feed 104.

The desulfurization unit 106 can operate at a temperature between about300° C. and about 450° C., such as about 300° C., about 350° C., about400° C., about 450° C., or another temperature. The desulfurization unit106 can operate at a pressure between about 30 bar and about 180 bar,such as about 30 bar, about 60 bar, about 90 bar, about 120 bar, about150 bar, about 180 bar, or another pressure.

The desulfurized stream 108 from the desulfurization unit 106 has steam110 added before being reformed in a pre-reformer 112, for example, witha Ni based catalyst, among others. As used herein, a pre-reformergenerally converts a hydrocarbon stream to a methane rich streamcontaining CH₄, H₂, CO, CO₂ and steam. A reformer is generallydesignated to full conversion of a hydrocarbon stream to a synthesis gascontaining mainly H₂, CO, CO₂, steam, and a small amount of CH₄.

In the pre-reformer 112, the hydrocarbons are reacted with the steam110. The pre-reformer 112 can operate at a pressure between about 0.01bar and about 50 bar, such as about 0.01 bar, about 0.1 bar, about 0.5bar, about 1 bar, about 5 bar, about 10 bar, about 20 bar, about 30 bar,about 40 bar, about 50 bar, or another pressure. The molar ratio ofhydrogen to hydrocarbon at the outlet of the pre-reformer 112 can bebetween about 1:1 and about 10:1, such as about 1:1, about 2:1, about4:1, about 6:1, about 8:1, about 10:1, or another ratio. Thepre-reformer 112 can operate at a temperature between about 300° C. andabout 550° C., such as about 300° C., about 400° C., about 450° C.,about 500° C., about 550° C., or another temperature.

The pre-reformer 112 produces a methane rich gas, or reformate 114, inwhich at least almost all of the C₂ ⁺ hydrocarbons have been convertedto C₁. The reformate 114 from the pre-reformer 112 is passed to thehydrogen membrane separator 102. The hydrogen membrane separator 102utilizes a hydrogen selective membrane that operates at hightemperatures. In some embodiments, the hydrogen membrane separator 102operates at temperatures that range between about 300° C. and about 550°C. In some embodiments, the hydrogen membrane separator 102 comprises aproton conducting material, which is electrically driven to transportthe hydrogen to the permeation side.

In various embodiments, the hydrogen selective membrane includespalladium, palladium alloy, carbon based membranes, or zeolite basedmembranes. The membrane selectively removes hydrogen, in a permeatestream 116, from the reformate 114. The selection of the membrane may bebased on cost, and other factors, such as ease of manufacturing,lifespan, and hydrogen flux.

The hydrogen is further treated, for example, if needed to reach apurity specification for fuel cell applications, such as in vehicles.According to ISO FDIS 1467-2, hydrogen purity is generally, at least,99.97%. For example, a pressure swing adsorption (PSA) system (notshown) may be included in the trigeneration facility 100 for thepurification of hydrogen. In these embodiments, the permeate stream 116is routed to the PSA system. The PSA system may include two columnsfilled with a zeolite absorbent, one active column and one regeneratingcolumn. The permeate stream 116 is flowed through the active column,which absorbs impurities from the hydrogen flow. In embodiments, thepurity of the hydrogen in the permeate stream 116 is greater than about80 vol. %, greater than about 90 vol. %, greater than about 95 vol. %,or higher. Once impurities start to break through the active column, theflow is switched over to the regenerating column, which then becomes theactive column. The previously active column is then regenerated.

After purification, the hydrogen is compressed to 400 to 900 bar asneeded for dispensing to fuel cell vehicles. The flow of hydrogen in thepermeate stream 116 can be increased through controlling the permeateside pressure or the membrane surface area. As the pressure of 114 isincreased, hydrogen flux through the membrane will increase resulting ina higher flow of the permeate stream 116. Alternatively, a higherpressure of reformate 114 coupled with a higher surface area, would alsoresult in a higher flow of the permeate stream 116.

The retentate stream 118 from the membrane separator is routed to ananode side of a solid oxide fuel cell (SOFC) 124. In some embodiments, aheat exchanger 120 is used to adjust the temperature of the retentatestream 118 forming a temperature-controlled stream 122. The temperatureof the temperature-controlled stream 122 may be increased to minimizethe temperature gradient between the anode and cathode sides, as well asavoiding thermal shock (since the SOFC operates at a much highertemperature). Air, or another oxidizer, is fed to the cathode side ofthe SOFC 124 and the electrochemical reaction produces electricity 126.Hydrocarbons in the retentate 118 may be further reformed in the SOFC124 increasing the amount of synthesis gas in the anode exhaust 128. Theanode exhaust 128 can be recycled back to the membrane reactor or usedto produce heat 130 for preheating the feed streams, such as by flowingthrough the heat exchanger 120 to increase the temperature of thereformate 118 and form the temperature-controlled stream 122. As thegeneration of electricity in the SOFC 124 is exothermic, othertechniques can be used to remove heat 130 from the SOFC 124. Forexample, more air in passed through the cathode side of the SOFC 124, orthe SOFC 124 may be enclosed in a heat exchanger, or include coolingcoils within the SOFC 124 structure , among other techniques.Accordingly, the trigeneration facility 100 allows for the optimizationof the production of hydrogen in the permeate stream 116, electricity126, and heat 130

FIG. 2 is a schematic drawing of a trigeneration facility 200 thatincorporates a hydrogen membrane separator with an integrated water-gasshift reactor. Similar numbered items are as described with respect toFIG. 1. Further, the trigeneration facility 200 of FIG. 2 may alsoinclude a PSA system for the purification of hydrogen in the permeate116.

In this example, the reformate 114 is passed to a membrane reactor 202in which a water-gas shift catalyst is packed along with a hightemperature membrane. In some embodiments, the water-gas shift catalystmay be a high temperature iron-oxide based or intermediate-lowtemperature copper oxide based catalysts. The presence of the water-gasshift catalyst allows the conversion of carbon monoxide in the reformate114 to hydrogen and CO2.

FIG. 3 is a flowchart of a method 300 for using a hydrocarbon feedstream to produce hydrogen, electricity, and heat in a trigenerationfacility that incorporates a hydrogen membrane separator. The method 300begins at block 302, with the desulfurization of the hydrocarbon feedstream. As described herein, this may be performed using adesulfurization catalyst or, if the amount of sulfur is low, usingadsorbents. In various embodiments, the adsorbents include activatedcarbon, CuO/ZnO/Al₂O₃, Ag-zeolite, or Ag/CeO₂, among others. Theadsorbents do not need to be operated at high temperatures, but may beoperated at ambient temperatures, such as between about 5° C. and about50° C. Selection of the adsorbents is dependent on the sulfur compoundsand type of hydrocarbon feedstock. For example, different adsorbents maybe used for H₂S, COS, heteroatom hydrocarbons that contain sulfur, andthe like. The low temperature techniques are economical for small scaleand simplified maintenance, e.g. replacing a cartridge of sulfuradsorbents in a passenger vehicle.

At block 304, the desulfurized feed stream is reformed, for example, ina pre-reformer. Steam is added to the desulfurized feed stream upstreamof the pre-reformer.

At block 306, hydrogen is separated from the reformate in a membraneseparator as a permeate stream. As described herein, the membraneseparator may be run at an elevated temperature.

At block 308, a retentate stream from the membrane separator is providedto the anode of an SOFC. In the SOFC, electricity is generated from theretentate stream and an oxidizer stream provided to the cathode. Aportion of the anode exhaust stream may be recycled back to be blendedwith the desulfurized feed stream, at block 304, supplying the requiredsteam to the pre-reformer or may be processed for sequestration ofcarbon dioxide or may be processed to recover the remaining fuel valuefor heat.

At block 310, the permeate stream may be processed to purify thehydrogen, for example for use in transportation applications. Thepurification may be performed in a pressure swing adsorption, amongothers.

At block 312, process heat from the SOFC is provided to other processunits, such as a heat exchanger used to increase the temperature of theretentate stream prior to providing the retentate stream to the anode ofthe SOFC. In other embodiments, the process heat is used to generatesteam for chemical plant.

FIG. 4 is a flowchart of a method 400 for using a feed stream to producehydrogen, electricity, and heat in a trigeneration facility thatincorporates a hydrogen membrane separator with an integrated water-gasshift reactor. Like numbered items are as described with respect to FIG.3. In this method 400, at block 402, a water-gas shift reaction isperformed on the reformate prior to or during the separation of thehydrogen. In some embodiments, a water-gas shift catalyst is included inthe membrane separator to increase the yield of hydrogen and CO₂.

EXAMPLES Process Simulations

Process models were built using Aspen plus (version 10) for theembodiments described with respect to FIGS. 1 and 2, and all thecomponents of the process are modeled to determine the mass and energyflows. These process models contain detailed routing of various streamsthat can satisfy the pre-heating or cooling of process streams bymatching the available heat contents of hot streams with pre-heatingheat load requirements of components of the process. They also representhow the hydrogen selective membrane in the hydrogen membrane separators102 and 202 and the SOFC 124 can be represented on Aspen plus modelingplatform as these embodiments are not standard, “plug-and-play” modelingmodules in Aspen plus library.

FIG. 5 is a simplified process model 500 of a trigeneration facilitythat incorporates a hydrogen membrane separator. The results of themodeling for S/C=3 are shown in Tables 1 to 3. The results of themodeling for S/C=4 are shown in Tables 4 to 6. S/C is a steam to carbonratio and defined as the molar flowrate of steam over the molar flowrateof hydrocarbon feed times its carbon number; for example for a givenS/C=3, the molar flowrate of steam divided by the sum of the molarflowrate propane times 3 and the molar flowrate of butane times 4 equals3.

TABLE 1 Pre-reformer operation at S/C = 3 Units PRE-IN MEMBR-INTemperature C. 563.06 494.23 Pressure bar 10.00 10.00 Mole Flows WATERmol/min 137.75 109.58 H2 mol/min 0.64 24.77 O2 mol/min 0.00 0.00 N2mol/min 0.00 0.00 CH4 mol/min 0.00 31.54 CO2 mol/min 0.00 13.80 COmol/min 0.00 0.58 PROPANE mol/min 6.56 0.00 BUTANE mol/min 6.56 0.00Mole Fractions WATER 91%  61%  H2 0% 14%  O2 0% 0% N2 0% 0% CH4 0% 17% CO2 0% 8% CO 0% 0% PROPANE 4% 0% BUTANE 4% 0%

TABLE 2 Membrane operation at S/C = 3 Units MEMBR-IN RETENTAT DEPRERETH2PERM H2PERM2 Temperature ° C. 494.23 494.23 494.23 494.23 494.23Pressure bar 10.00 10.00 1.05 10.00 1.02 Mole Flows WATER mol/min 109.58109.58 109.58 0.00 0.00 H2 mol/min 24.77 4.95 4.95 19.82 19.82 O2mol/min 0.00 0.00 0.00 0.00 0.00 N2 mol/min 0.00 0.00 0.00 0.00 0.00 CH4mol/min 31.54 31.54 31.54 0.00 0.00 CO2 mol/min 13.80 13.80 13.80 0.000.00 CO mol/min 0.58 0.58 0.58 0.00 0.00 PROPANE mol/min 0.00 0.00 0.000.00 0.00 BUTANE mol/min 0.00 0.00 0.00 0.00 0.00 Mole Fractions WATER61% 68% 68%  0%   0% H2 14%  3%  3% 100% 100% O2  0%  0%  0%  0%   0% N2 0%  0%  0%  0%   0% CH4 17% 20% 20%  0%   0% CO2  8%  9%  9%  0%   0%CO  0%  0%  0%  0%  0% PROPANE  0%  0%  0%  0%  0% BUTANE  0%  0%  0% 0%  0%

TABLE 3 Solid Oxide Fuel Cell operation at S/C = 3 Units ANO-IN ANO-OUTTemperature ° C. 421.64 750.00 Pressure bar 1.05 1.05 Mole Flows WATERmol/min 109.58 161.15 H2 mol/min 4.95 16.46 O2 mol/min 0.00 0.00 N2mol/min 0.00 0.00 CH4 mol/min 31.54 0.00 CO2 mol/min 13.80 42.63 COmol/min 0.58 3.29 PROPANE mol/min 0.00 0.00 BUTANE mol/min 0.00 0.00Mole Fractions WATER 68%  72%  H2 3% 7% O2 0% 0% N2 0% 0% CH4 20%  0%CO2 9% 19%  CO 0% 1% PROPANE 0% 0% BUTANE 0% 0%

TABLE 4 Pre-reformer operation at S/C = 4 Units PRE-IN MEMBR-INTemperature ° C. 579.50 492.69 Pressure bar 10.00 10.00 Mole Flows WATERmol/min 183.67 152.29 H2 mol/min 0.64 31.18 O2 mol/min 0.00 0.00 N2mol/min 0.00 0.00 CH4 mol/min 0.00 29.94 CO2 mol/min 0.00 15.40 COmol/min 0.00 0.58 PROPANE mol/min 6.56 0.00 BUTANE mol/min 6.56 0.00Mole Fractions WATER 93%  66%  H2 0% 14%  O2 0% 0% N2 0% 0% CH4 0% 13% CO2 0% 7% CO 0% 0% PROPANE 3% 0% BUTANE 3% 0%

TABLE 5 Membrane operation at S/C = 4 Units MEMBR-IN RETENTAT DEPRERETH2PERM H2PERM2 Temperature ° C. 492.69 492.69 492.69 492.69 492.69Pressure bar 10.00 10.00 1.05 10.00 1.02 Mole Flows WATER mol/min 152.29152.29 152.29 0.00 0.00 H2 mol/min 31.18 6.24 6.24 24.94 24.94 O2mol/min 0.00 0.00 0.00 0.00 0.00 N2 mol/min 0.00 0.00 0.00 0.00 0.00 CH4mol/min 29.94 29.94 29.94 0.00 0.00 CO2 mol/min 15.40 15.40 15.40 0.000.00 CO mol/min 0.58 0.58 0.58 0.00 0.00 PROPANE mol/min 0.00 0.00 0.000.00 0.00 BUTANE mol/min 0.00 0.00 0.00 0.00 0.00 Mole Fractions WATER66% 74% 74%  0%  0% H2 14%  3%  3% 100% 100% O2  0%  0%  0%  0%  0% N2 0%  0%  0%  0%  0% CH4 13% 15% 15%  0%  0% CO2  7%  8%  8%  0%  0% CO 0%  0%  0%  0%  0% PROPANE  0%  0%  0%  0%  0% BUTANE  0%  0%  0%  0% 0%

TABLE 6 Solid Oxide Fuel Cell operation at S/C = 4 Units ANO-IN ANO-OUTTemperature C. 434.20 750.00 Pressure bar 1.05 1.05 Mole Flows WATERmol/min 152.29 202.07 H2 mol/min 6.24 16.34 O2 mol/min 0.00 0.00 N2mol/min 0.00 0.00 CH4 mol/min 29.94 0.00 CO2 mol/min 15.40 43.27 COmol/min 0.58 2.65 PROPANE mol/min 0.00 0.00 BUTANE mol/min 0.00 0.00Mole Fractions WATER 74%  76%  H2 3% 6% O2 0% 0% N2 0% 0% CH4 15%  0%CO2 8% 16%  CO 0% 1% PROPANE 0% 0% BUTANE 0% 0%

FIG. 6 is a simplified process model of a trigeneration facility thatincorporates a hydrogen membrane separator with an integrated water gasshift reactor. In the simulations based on the trigeneration facility ofFIG. 6, the pre-reformer is run at a temperature of 493° C. and apressure of 10 bar. The results of the modeling of the system for S/C=3with a water-gas shift reaction present before the hydrogen separationmembrane are shown in Tables 7 to 12.

TABLE 7 Products after pre-reformer (without H2O) at S/C = 3 PRO- BU- H2O2 N2 CH4 CO2 CO PANE TANE 35.04% 0.00% 0.00% 44.62% 19.52% 0.81% 0.00%0.00%

TABLE 8 Operation of pre-reformer at S/C = 3 Units PRE-IN AFT-PRETemperature C. 563.06 494.23 Pressure bar 10.00 10.00 Mole Flows WATERmol/min 137.75 109.58 H2 mol/min 0.64 24.77 O2 mol/min 0.00 0.00 N2mol/min 0.00 0.00 CH4 mol/min 0.00 31.54 CO2 mol/min 0.00 13.80 COmol/min 0.00 0.58 PROPANE mol/min 6.56 0.00 BUTANE mol/min 6.56 0.00Mole Fractions WATER 91%  61%  H2 0% 14%  O2 0% 0% N2 0% 0% CH4 0% 17% CO2 0% 8% CO 0% 0% PROPANE 4% 0% BUTANE 4% 0%

TABLE 9 Operation of water gas shift unit at S/C = 3 Units COOL-PREAFT-WGS Temperature C. 400.00 402.95 Pressure bar 10.00 10.00 Mole FlowsWATER mol/min 109.58 109.01 H2 mol/min 24.77 25.33 O2 mol/min 0.00 0.00N2 mol/min 0.00 0.00 CH4 mol/min 31.54 31.54 CO2 mol/min 13.80 14.36 COmol/min 0.58 0.01 PROPANE mol/min 0.00 0.00 BUTANE mol/min 0.00 0.00Mole Fractions WATER 61%  60%  H2 14%  14%  O2 0% 0% N2 0% 0% CH4 17% 17%  CO2 8% 8% CO 0% 0% PROPANE 0% 0% BUTANE 0% 0%

TABLE 10 Operation of membrane at S/C = 3 Units AFT-WGS RETENTATDEPRERET H2PERM H2PERM2 Temperature C. 402.95 402.95 402.95 402.95402.95 Pressure bar 10.00 10.00 1.05 10.00 1.02 Mole Flows WATER mol/min109.01 109.01 109.01 0.00 0.00 H2 mol/min 25.33 5.07 5.07 20.27 20.27 O2mol/min 0.00 0.00 0.00 0.00 0.00 N2 mol/min 0.00 0.00 0.00 0.00 0.00 CH4mol/min 31.54 31.54 31.54 0.00 0.00 CO2 mol/min 14.36 14.36 14.36 0.000.00 CO mol/min 0.01 0.01 0.01 0.00 0.00 PROPANE mol/min 0.00 0.00 0.000.00 0.00 BUTANE mol/min 0.00 0.00 0.00 0.00 0.00 Mole Fractions WATER60% 68% 68%  0%  0% H2 14%  3%  3% 100% 100% O2  0%  0%  0%  0%  0% N2 0%  0%  0%  0%  0% CH4 17% 20% 20%  0%  0% CO2  8%  9%  9%  0%  0% CO 0%  0%  0%  0%  0% PROPANE  0%  0%  0%  0%  0% BUTANE  0%  0%  0%  0% 0%

TABLE 11 Operation of solid oxide fuel cell at S/C = 3 Units ANO-INANO-OUT Temperature C. 400.00 750.00 Pressure bar 1.05 1.05 Mole FlowsWATER mol/min 109.01 160.77 H2 mol/min 5.07 16.40 O2 mol/min 0.00 0.00N2 mol/min 0.00 0.00 CH4 mol/min 31.54 0.00 CO2 mol/min 14.36 42.63 COmol/min 0.01 3.29 PROPANE mol/min 0.00 0.00 BUTANE mol/min 0.00 0.00Mole Fractions WATER 68%  72%  H2 3% 7% O2 0% 0% N2 0% 0% CH4 20%  0%CO2 9% 19%  CO 0% 1% PROPANE 0% 0% BUTANE 0% 0%

TABLE 12 Productivity of system at S/C = 3 SOFC H2 Stack SYSTEM Prod.ELEC PROPANE BUTANE POWER SOFCV SOFCEFF REQPOW SOFCELC H2 Prod EFF EffEFF FU* (MOL/MIN) (MOL/MIN) (KW) (VOLT) (%) (KW) (KW) (KG/DAY) (%) (%)(%) 0.85 6.56 6.56 255.40 0.71 65.00 23.14 224.60 56.98 57.09 28.1248.44 *FU is fuel utilization.

The results of the modeling of the system for S/C=4 with a water-gasshift reaction present before the hydrogen separation membrane are shownin Tables 13 to 18.

TABLE 13 Products after pre-reformer (without H2O) at S/C = 4 PRO- BU-H2 O2 N2 CH4 CO2 CO PANE TANE 40.44% 0.00% 0.00% 38.84% 19.98% 0.75%0.00% 0.00%

TABLE 14 Operation of pre-reformer at S/C = 4 Units PRE-IN AFT-PRETemperature C. 579.50 492.69 Pressure bar 10.00 10.00 Mole Flows WATERmol/min 183.67 152.29 H2 mol/min 0.64 31.18 O2 mol/min 0.00 0.00 N2mol/min 0.00 0.00 CH4 mol/min 0.00 29.94 CO2 mol/min 0.00 15.40 COmol/min 0.00 0.58 PROPANE mol/min 6.56 0.00 BUTANE mol/min 6.56 0.00Mole Fractions WATER 93%  66%  H2 0% 14%  O2 0% 0% N2 0% 0% CH4 0% 13% CO2 0% 7% CO 0% 0% PROPANE 3% 0% BUTANE 3% 0%

TABLE 15 Operation of water gas shift unit at S/C = 4 Units COOL-PREAFT-WGS Temperature C. 400.00 402.37 Pressure bar 10.00 10.00 Mole FlowsWATER mol/min 152.29 151.73 H2 mol/min 31.18 31.74 O2 mol/min 0.00 0.00N2 mol/min 0.00 0.00 CH4 mol/min 29.94 29.94 CO2 mol/min 15.40 15.97 COmol/min 0.58 0.01 PROPANE mol/min 0.00 0.00 BUTANE mol/min 0.00 0.00Mole Fractions WATER 66%  66%  H2 14%  14%  O2 0% 0% N2 0% 0% CH4 13% 13%  CO2 7% 7% CO 0% 0% PROPANE 0% 0% BUTANE 0% 0%

TABLE 16 Operation of membrane at S/C = 4 Units AFT-WGS RETENTATDEPRERET H2PERM H2PERM2 Temperature C. 402.37 402.37 402.37 402.37402.37 Pressure bar 10.00 10.00 1.05 10.00 1.02 Mole Flows WATER mol/min151.73 151.73 151.73 0.00 0.00 H2 mol/min 31.74 6.35 6.35 25.39 25.39 O2mol/min 0.00 0.00 0.00 0.00 0.00 N2 mol/min 0.00 0.00 0.00 0.00 0.00 CH4mol/min 29.94 29.94 29.94 0.00 0.00 CO2 mol/min 15.97 15.97 15.97 0.000.00 CO mol/min 0.01 0.01 0.01 0.00 0.00 PROPANE mol/min 0.00 0.00 0.000.00 0.00 BUTANE mol/min 0.00 0.00 0.00 0.00 0.00 Mole Fractions WATER66% 74% 74%   0%   0% H2 14%  3%  3% 100% 100% O2  0%  0%  0%   0%   0%N2  0%  0%  0%   0%   0% CH4 13% 15% 15%   0%   0% CO2  7%  8%  8%   0%  0% CO  0%  0%  0%   0%   0% PROPANE  0%  0%  0%   0%   0%

TABLE 17 Operation of solid oxide fuel cell at S/C = 4 Units ANO-INANO-OUT Temperature C. 400.00 750.00 Pressure bar 1.05 1.05 Mole FlowsWATER mol/min 151.73 201.68 H2 mol/min 6.35 16.28 O2 mol/min 0.00 0.00N2 mol/min 0.00 0.00 CH4 mol/min 29.94 0.00 CO2 mol/min 15.97 43.28 COmol/min 0.01 2.64 PROPANE mol/min 0.00 0.00 BUTANE mol/min 0.00 0.00Mole Fractions WATER 74%  76%  H2 3% 6% O2 0% 0% N2 0% 0% CH4 15%  0%CO2 8% 16%  CO 0% 1% PROPANE 0% 0% BUTANE 0% 0%

TABLE 18 Productivity of system at S/C = 4 SOFC H2 Stack Prod ELECPROPANE BUTANE POWER SOFC V SOFCEFF REQPOW SOFCELC H2 Prod SYSTEM EffEFF FU* (MOL/MIN) (MOL/MIN) (KW) (VOLT) (%) (KW) (KW) (KG/DAY) EFF (%)(%) (%) 0.85 6.56 6.56 247.04 0.72 65.00 27.06 212.57 71.85 59.32 34.2348.39 *FU is the fuel utilization.

Exemplary Embodiments

An embodiment described herein provides a method for coproduction ofhydrogen, electrical power, and heat energy. The method includesdesulfurizing a feed stream to form a desulfurized feed stream,reforming the desulfurized feed stream to form a methane rich gas, andproviding the methane rich gas to a membrane separator. A hydrogenstream is produced in a permeate from the membrane separator. Aretentate stream from the membrane separator is provided to a solidoxide fuel cell (SOFC). Electrical power is produced in the SOFC fromthe retentate stream.

In an aspect, the method includes blending hydrogen with the feed streamprior to desulfurizing the feed stream. In an aspect, a portion of thehydrogen stream is mixed with the feed stream prior to desulfurizing thefeed stream. In an aspect, the method includes desulfurizing the feedstream in an adsorption unit.

In an aspect, the method includes compressing the hydrogen stream foruse as a product. In an aspect, the method includes heating theretentate stream to an operating temperature for the SOFC prior toproviding the retentate stream to the SOFC.

In an aspect, the method includes utilizing heat produced in the SOFC.In an aspect, the method includes heating the retentate stream with heatproduced in the SOFC. In an aspect, the method includes generating steamwith the heat produced in the SOFC.

Another embodiment described herein provides a trigeneration facility.The trigeneration facility includes a desulfurization unit to removesulfur from a hydrocarbon feed stream forming a desulfurized feedstream. A pre-reformer is included to convert the desulfurized feedstream to a methane rich gas. A membrane separator is included to removeat least a portion of hydrogen from the methane rich gas in a permeate.A solid oxide fuel cell (SOFC) is included to generate electrical powerfrom a retentate from the membrane separator.

In an aspect, the trigeneration facility includes a heat exchanger toutilize heat energy from the SOFC. In an aspect, the trigenerationfacility includes a heat exchanger to heat the retentate to an operatingtemperature of the SOFC.

In an aspect, the pre-reformer includes a steam reformer including anickel catalyst. In an aspect, the pre-reformer operates at atemperature between about 300° C. and about 550° C. in an aspect, thesteam reformer operates at a pressure of between about 2 bar and about30 bar.

In an aspect, the trigeneration facility includes a water-gas shiftcatalyst to increase the amount of hydrogen in the methane rich gas. Inan aspect, the water-gas shift catalyst includes iron oxides or copperoxides.

In an aspect, the membrane separator is a high-temperature hydrogenselective membrane. In an aspect, the membrane separator includespalladium, or a palladium alloy, or both. In an aspect, the membraneseparator includes a carbon-based membrane, or a zeolite based membrane,or both. In an aspect, the membrane separator includes an integratedwater-gas shift catalyst. In an aspect, the membrane separator includesa proton conducting material, which is be electrically driven totransport the hydrogen to the permeation side.

In an aspect, the feed stream includes propane, or butane, or both. Inan aspect, the feed stream includes liquefied natural gas. In an aspect,the feed stream includes a raw natural gas.

Other implementations are also within the scope of the following claims.

What is claimed is:
 1. A method for coproduction of hydrogen, electricalpower, and heat energy, comprising: desulfurizing a feed stream to forma desulfurized feed stream; reforming the desulfurized feed stream toform a methane rich gas; providing the methane rich gas to a membraneseparator; producing a hydrogen stream in a permeate from the membraneseparator; providing a retentate stream from the membrane separator to asolid oxide fuel cell (SOFC); and producing electrical power in the SOFCfrom the retentate stream.
 2. The method of claim 1, comprising blendinghydrogen with the feed stream prior to desulfurizing the feed stream. 3.The method of claim 1, comprising mixing a portion of the hydrogenstream with the feed stream prior to desulfurizing the feed stream. 4.The method of claim 1, comprising desulfurizing the feed stream in aadsorption unit.
 5. The method of claim 1, comprising compressing thehydrogen stream for use as a product.
 6. The method of claim 1,comprising heating the retentate stream to an operating temperature forthe SOFC prior to providing the retentate stream to the SOFC.
 7. Themethod of claim 1, comprising utilizing heat produced in the SOFC. 8.The method of claim 7, comprising heating the retentate stream with heatproduced in the SOFC.
 9. The method of claim 7, comprising generatingsteam with the heat produced in the SOFC.
 10. A trigeneration facility,comprising: a desulfurization unit to remove sulfur from a hydrocarbonfeed stream forming a desulfurized feed stream; a pre-reformer toconvert the desulfurized feed stream to a methane rich gas; a membraneseparator to remove at least a portion of hydrogen from the methane richgas in a permeate; and a solid oxide fuel cell (SOFC) to generateelectrical power from a retentate from the membrane separator.
 11. Thetrigeneration facility of claim 10, comprising a heat exchanger toutilize heat energy from the SOFC.
 12. The trigeneration facility ofclaim 10, comprising a heat exchanger to heat the retentate to anoperating temperature of the SOFC.
 13. The trigeneration facility ofclaim 10, wherein the pre-reformer comprises a steam reformer comprisinga nickel catalyst.
 14. The trigeneration facility of claim 13, whereinthe pre-reformer operates at a temperature between about 300° C. andabout 550° C.
 15. The trigeneration facility of claim 13, wherein thesteam reformer operates at a pressure of between about 2 bar and about30 bar.
 16. The trigeneration facility of claim 10, comprising awater-gas shift catalyst to increase an amount of hydrogen in themethane rich gas.
 17. The trigeneration facility of claim 16, whereinthe water-gas shift catalyst comprises iron oxides or copper oxides. 18.The trigeneration facility of claim 10, wherein the membrane separatoris a high-temperature hydrogen selective membrane.
 19. The trigenerationfacility of claim 18, wherein the membrane separator comprisespalladium, or a palladium alloy, or both.
 20. The trigeneration facilityof claim 18, wherein the membrane separator comprises a carbon-basedmembrane, or a zeolite based membrane, or both.
 21. The trigenerationfacility of claim 10, wherein the membrane separator comprises anintegrated water-gas shift catalyst.
 22. The trigeneration facility ofclaim 10, wherein the membrane separator comprises a proton conductingmaterial, which is electrically driven to transport the hydrogen to thepermeation side.
 23. The trigeneration facility of claim 10, wherein thefeed stream comprises propane, or butane, or both.
 24. The trigenerationfacility of claim 10, wherein the feed stream comprises liquefiednatural gas.
 25. The trigeneration facility of claim 10, wherein thefeed stream comprises a raw natural gas.